Integration of side riser for aromatization of light paraffins

ABSTRACT

Systems and methods are provided for conversion of light paraffinic gases to form liquid products in a two-stage reaction system. In a first stage, the light paraffinic gas is exposed to heat transfer particles in a side riser, where the heat transfer particles correspond to particles used in a separate process. Examples of a separate process include fluidized coking and fluid catalytic cracking. The conditions in the side riser can be selected to allow for conversion of at least a portion of the paraffins to olefins. After conversion, the converted olefin stream is passed to the second reaction stage while the heat transfer particles are returned to the separate process. The converted olefin stream is then exposed to a conversion catalyst under conditions for forming aromatics from the converted olefin stream in a second reaction stage. By performing the initial alkane conversion to olefins in the first reaction stage, the amount of coke formed during the aromatics formation process is reduced or minimized.

FIELD

This invention relates to reactors, associated reactor systems, and processes for conversion of paraffin-containing streams to aromatics by using excess heat and regenerated particles from a cracking process to perform alkane dehydrogenation in a side riser followed by aromatization in a separate reaction stage.

BACKGROUND

Liquefied Petroleum Gas (LPG) is a petroleum product that typically includes primarily a mixture of C₃ and C₄ paraffins. LPG can be formed from a variety of processes. At a production site, LPG can be formed as part of various separations schemes. For example, due to restrictions on the energy content of natural gas for pipeline transport, LPG can be separated from some natural gas streams at the production site. Other sources of LPG can include formation of LPG as a secondary product in various refining and/or chemical production processes. Although LPG has some fuel value, it would be desirable to have additional options for upgrading LPG to higher value products.

Some conventional options for conversion of LPG to higher value products can involve converting at least a portion of the LPG into aromatic products. While such methods can be at least partially effective, conventional methods for conversion of C₃-C₄ paraffins to aromatics often have difficulties with limited production run lengths due to excessive coke formation on the catalysts used for aromatics formation.

U.S. Pat. No. 4,912,273 describes a method for producing aromatic hydrocarbons from alkanes. In an initial step, alkanes are introduced into a catalyst cooler vessel associated with a fluid catalytic cracking system to convert a portion of the alkanes to olefins. The olefins are then passed into a separate reactor stage for conversion of olefins and alkanes to aromatics.

U.S. Patent Application Publication 2015/0158789 describes an integrated system and corresponding method for processing C₃+ gas phase products generated at an extraction site to convert the gas phase products to liquid products. The liquid products can then be transported for further processing along with the extracted liquid phase products.

SUMMARY OF THE INVENTION

In various aspects, a method for processing a paraffin-containing feed is provided. The method includes passing a first portion of heat transfer particles from a burner of a reaction system to a reactor of the reaction system. A second portion of heat transfer particles can be passed from the burner to a side riser. A feed comprising C₃₊ paraffins can be exposed to the second portion of heat transfer particles in the side riser under paraffin to olefin conversion conditions to form a side riser effluent comprising heat transfer particles including deposited coke and a gas phase effluent. A third portion of the heat transfer particles including deposited coke can be separated from the gas phase effluent. At least a portion of the gas phase effluent can be exposed to one or more beds of a conversion catalyst to form an aromatic formation effluent comprising C₆-C₁₂ aromatics. The third portion of the heat transfer particles can be passed into the reaction system. A second feedstock can be exposed to the first portion of heat transfer particles and the third portion of the heat transfer particles in the reactor under first processing conditions to form a first reactor effluent comprising heat transfer particles including additional coke. At least a portion of the heat transfer particles including additional coke can be separated from the first reactor effluent. The separated heat transfer particles can then be passed into the burner.

In some aspects, the reaction system can correspond to a fluidized coking reaction system, with the heat transfer particles corresponding to coke particles. In such aspects, the reactor can correspond to a fluidized coking reactor. Optionally, the burner can correspond to a gasifier. In aspects where the burner corresponds to a gasifier, the reaction system can optionally further include a heater, with the heater providing indirect fluid communication between the reactor and the gasifier for at least a portion of the heat transfer particles. In other aspects, the reaction system can correspond to a fluid catalytic cracking reaction system, with the heat transfer particles corresponding to fluid catalytic cracking catalyst particles. In such aspects, the reactor can correspond to a fluid catalytic cracking reactor and the burner can correspond to a regenerator.

In various aspects, a reaction system for conversion of paraffins to aromatics is also provided. The reaction system includes a first reaction system comprising a reactor and a burner, the reactor being in fluid communication with the burner for transfer of heat transfer particles. The reaction system further includes a side riser comprising one or more side riser inlets and a side riser outlet, the one or more side riser inlets being in fluid communication with a feed source and in fluid communication with the burner for transfer of heat transfer particles. The reaction system further includes a separation stage comprising a separation stage inlet, a separation stage solids outlet, and a separation stage gas outlet, the separation stage inlet being in fluid communication with the side riser outlet, the separation stage solids outlet being in fluid communication with the first reactor. Additionally, the reaction system includes a second reaction system comprising one or more beds of conversion catalyst, the second reaction system being in fluid communication with the separation stage gas outlet. The conversion catalyst can include a) 0.1 wt % to 5.0 wt % of a metal from Groups 3-13 of the periodic table relative to a weight of the conversion catalyst, the metal optionally comprising Ga, In, or a combination thereof, and b) ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, MCM-22, MCM-49, or a combination thereof.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows an example of a configuration for integrating a process for conversion of paraffins to aromatics with a separate process.

FIG. 2 shows another example of a configuration for integrating a process for conversion of paraffins to aromatics with a separate process.

FIG. 3 shows an example of a fluidized coking reaction system.

FIG. 4 shows another example of a fluidized coking reaction system.

DETAILED DESCRIPTION

All numerical values within the detailed description and the claims herein are modified by “about” or “approximately” the indicated value, and take into account experimental error and variations that would be expected by a person having ordinary skill in the art.

Overview

In various aspects, systems and methods are provided for conversion of light paraffinic gases to form liquid products in a two-stage reaction system. In a first stage, the light paraffinic gas is exposed to heat transfer particles in a side riser, where the heat transfer particles correspond to particles used in a separate process. The heat transfer particles can correspond to fluid catalytic cracking catalyst particles, coke particles, or another type of particle that is passed between a reactor and a regenerator/gasifier/burner for a separate process. Optionally, the heat transfer particles can optionally be passed between the reactor and the regenerator/gasifier/burner via another intermediate vessel, such as a “heater” vessel in a fluidized coking system that includes a coking reactor, a heater, and a gasifier. Examples of a separate process include fluidized coking and fluid catalytic cracking. The conditions in the side riser can be selected to allow for conversion of at least a portion of the paraffins to olefins. After conversion, the converted olefin stream is passed to a second reaction stage while the heat transfer particles are returned to the separate process. The converted olefin stream is then exposed to conditions for forming aromatics from the converted olefin stream in the second reaction stage. By performing the initial alkane conversion to olefins in the first reaction stage, the amount of coke formed during the aromatics formation process is reduced or minimized. Instead, the coke associated with the initial conversion of paraffins to olefins is formed in the first stage, such as forming coke on the heat transfer particles. The heat transfer particles can (eventually) be regenerated to remove the coke in the regenerator associated with the separate process.

Various processes within a refinery or chemical plant setting can generate minor portions C₃ and/or C₄ alkanes as part of a product slate. Traditionally, unless the C₄ paraffins have an unusually high concentration of isobutane, such C₃-C₄ paraffins generated in a refinery have ended up being used for fuel value. This is due to difficulties in finding a cost-effective method for upgrading such paraffins to higher value products.

In addition to paraffins generated within a refinery, many petroleum streams generated at a wellhead or other petroleum production site can include substantial quantities of C₃-C₄ paraffins. Prior to pipelining natural gas produced at a wellhead, it may be necessary to separate out C₃-C₄ paraffins from the natural gas in order to meet pipeline specifications. For crude fractions delivered to a refinery, any C₃-C₄ paraffins entrained in the crude fraction will likely be separated out during an initial fractionation. Traditionally, finding a high value use for these additional sources of C₃-C₄ paraffins has also been difficult.

One potential use for light paraffins is to convert the light paraffins to higher value olefins and/or aromatics. However, processes for conversion of alkanes (i.e., paraffins) to olefins and/or aromatics tend to result in substantial amounts of coke formation on the conversion catalyst. This rapid coking results in catalyst deactivation on a time scale that can quickly render the catalyst ineffective for performing further conversion. The rapid coking can be mitigated by performing the conversion in a reactor that allows for continuous regeneration of the conversion catalyst, such as a moving bed reactor. However, even with catalyst regeneration, typical catalysts for conversion of alkanes to aromatics can experience significant deactivation after a few regeneration cycles. Thus, it would be preferable to reduce or minimize the coke formed on the conversion catalyst. Additionally, due to the highly endothermic nature of the conversion of light paraffins to olefins and/or aromatics, having a separate process train for conversion of light paraffins can potentially be cost prohibitive relative to the final value of the converted products.

In various aspects, the difficulties with catalyst coke formation during a process for conversion of alkanes to aromatics can be reduced or minimized by performing an initial step of converting the alkanes to olefins in a separate process stage. The conversion of alkanes to olefins can be performed in a side riser or other reactor volume, where the side riser is in partial fluid communication with a separate process. The partial fluid communication can allow heat transfer particles from the separate process to enter the side riser while substantially avoiding transfer of process fluids to or from the separate process.

By using heat transfer particles from the separate process, at least a portion of the alkanes can be converted to olefins without being in the presence of the conversion catalyst used for aromatics formation. Instead, the conversion of alkanes to olefins is performed in the presence of the heat transfer particles. This means that the coke formed during formation of olefins from the alkanes is deposited on the heat transfer particles rather than on the subsequent conversion catalyst that is used for forming aromatics. This can substantially reduce the net amount of coke formed on the aromatic formation catalyst, thus providing a substantial extension in run length for the overall process. After forming olefins in the presence of the heat transfer particles, the heat transfer particles can be returned to the separate process while the olefin-containing effluent can be passed to the conversion catalyst for aromatic formation.

Although the heat transfer particles are distinct from the conversion catalyst for aromatic formation, the heat transfer particles may optionally also have catalytic activity that can assist with conversion of alkanes to olefins. The nature of this catalytic activity can be dependent on the source of the heat transfer particles. For example, in aspects where the heat transfer particles correspond to coke from a fluidized coking system (such as a Flexicoking system), the coke particles can include metals such as vanadium or nickel. Without being bound by any particular theory, it is noted that these metals within the coke particles can catalyze dehydrogenation chemistry, and therefore may facilitate dehydrogenation of alkanes to olefins. As another example, the acidic sites on fluid catalytic cracking (FCC) catalysts can catalyze cracking reactions. Without being bound by any particular theory, when FCC catalyst particles are used as the heat transfer particles, the acidic sites on the FCC catalyst particles may facilitate cracking of the feed, which could result in formation of a mixture of olefins with lower average molecular weight.

Definitions

The term “aromatic hydrocarbons” refers to molecules containing one or more aromatic rings. Examples of aromatic hydrocarbons are benzene, toluene, xylenes, naphthalene, and methylnaphthalenes.

The term “aromatic” refers to unsaturated compounds with at least one closed ring of at least 6 atoms, with all of the ring atoms being co-planar or almost co-planar and covalently linked, and with all of the ring atoms being part of a mesomeric system. As used herein, when the “aromatic” substituent is monocyclic, it preferably contains 6 ring atoms, and when the “aromatic” substituent is polycyclic, it preferably contains 10 ring atoms contained in fused rings.

The term “C_(n)” hydrocarbon refers to a hydrocarbon with “n” carbon atoms, and “C_(n)-C_(m) hydrocarbons” represents hydrocarbons having between “n” and “m” carbon atoms.

The term “catalyst” refers to a material, which under certain conditions of temperature or pressure increases the rate of specific chemical reactions. A catalyst may also be a material that performs as a physisorbent or chemisorbent for specific components of the feed stream.

The term “chain length” may broadly refer to a number of atoms forming and/or making a backbone and/or structure of a molecule and/or compound, such as carbon atoms for a hydrocarbon.

The term “chemical reaction” refers to any process including the breaking or making of chemical bonds including a dissociation, recombination, or rearrangement of atoms.

The term “coke” refers to the solid residue remaining from the pyrolysis of hydrocarbons.

The term “crude oil” refers to hydrocarbons formed primarily of carbon and hydrogen atoms. The hydrocarbons may also include other elements, such as, but not limited to, halogens, metallic elements, nitrogen, oxygen, or sulfur. Hydrocarbons derived from an oil-bearing formation may include, but are not limited to, kerogen, bitumen, pyrobitumen, asphaltenes, resins, oils, or combinations thereof.

The term “fixed-bed reactor” refers to a reactor containing catalyst material typically in pellet form, packed in a static bed. Such fixed-bed reactors can correspond to any convenient type of fixed-bed reactor, such as a radial flow reactor and/or an axial flow reactor.

The term “higher hydrocarbons” refers to hydrocarbon(s) having more than one carbon atom per molecule, e.g., ethane, ethylene, propane, propylene, benzene, toluene, xylenes, naphthalene, and/or methyl naphthalene; and/or organic compound(s) including at least one carbon atom and at least one non-hydrogen atom, e.g., methanol, ethanol, methylamine, and/or ethylamine.

The term “hydrocarbon” refers to an organic compound that includes primarily, if not exclusively, the elements hydrogen and carbon. Hydrocarbons may also include other elements, such as, but not limited to, halogens, metallic elements, nitrogen, oxygen, and/or sulfur. Hydrocarbons generally fall into two classes: aliphatic, or straight chain hydrocarbons, and cyclic, or closed ring hydrocarbons, including cyclic terpenes. Examples of hydrocarbon-containing materials include any form of natural gas, oil, coal, and bitumen.

The term “hydrocarbon diluent” refers to any substance containing one or more hydrocarbon compounds and/or substituted hydrocarbon compounds, which is suitable for use for diluting a hydrocarbon in the practice of the invention. For example, a tail gas stream containing hydrocarbons may be an added diluent for natural gas.

The term “hydrocarbon stream” refers to a hydrocarbon or mixtures of hydrocarbons that are gases or liquids. For example, hydrocarbon fluids may include a hydrocarbon or mixtures of hydrocarbons that are gases or liquids at formation conditions, at processing conditions or at ambient conditions (15° C. and 1 atm pressure). Hydrocarbon fluids may include, for example, oil, natural gas, coalbed methane, shale oil, pyrolysis oil, pyrolysis gas, a pyrolysis product of coal, and other hydrocarbons that are in a gaseous or liquid state

The term “light hydrocarbons” refer to hydrocarbons having carbon numbers in a range from 1 to 5.

The term “natural gas” refers to a multi-component gas obtained from a crude oil well (associated gas) or from a subterranean gas-bearing formation (non-associated gas). The composition and pressure of natural gas can vary significantly. A typical natural gas stream contains methane (C₁) as a significant component. Raw natural gas may also contain ethane (C₂), higher molecular weight hydrocarbons, acid gases (such as carbon dioxide, hydrogen sulfide, carbonyl sulfide, carbon disulfide, and mercaptans), and minor amounts of contaminants such as water, nitrogen, iron sulfide, wax, and crude oil. As used herein, natural gas includes gas resulting from the regasification of a liquefied natural gas, which has been purified to remove contaminates, such as water, acid gases, and most of the higher molecular weight hydrocarbons.

The term “high quality gas” refers to a gas that has undergone natural gas processing to separate various hydrocarbons and fluids from a raw natural gas. Also referred to as pipeline quality dry natural gas.

The term “raw natural gas” refers to a gas that is included of methane, but may also include numerous other light hydrocarbons including ethane, propane, and butanes. Higher molecular weight hydrocarbons, including pentanes, hexanes, and impurities like benzene may also be present in small amounts. Furthermore, raw natural gas may contain amounts of non-hydrocarbon impurities, such as nitrogen, hydrogen sulfide, carbon dioxide, and traces of helium, carbonyl sulfide, various mercaptans, and water.

In a conventional fluidized coking process, the fluidized coking system can include a coking reactor and a burner. The term “Flexicoking®” refers to a fluidized coking process that includes a coking reactor, a heater, and a gasifier. It is noted that a gasifier corresponds to a type of burner, even though the gasifier in a Flexicoking® process may be operated in a different manner than a conventional burner in a conventional fluidized coking process. It is further understood that Flexicoking® is type of fluidized coking, although the specific types of components in a conventional fluidized coking process may differ. In this discussion, references to fluidized coking are understood to include Flexicoking®, unless otherwise specified.

Conversion of Paraffins to Olefins in Side Riser

In various aspects, a side riser associated with a separate process can be used for conversion of paraffins to olefins. The side riser can receive a feed of light paraffins, such as C₃ and/or C₄ paraffins from a light paraffin source. The side riser can also receive a stream of heat transfer particles that are at an elevated temperature, such as a side stream of heat transfer particles withdrawn from a burner (such as a regenerator or gasifier) and/or a heater for the separate process.

The feed for the side riser can correspond to any convenient feed source (including combinations of streams from various sources) that includes C₃ and/or C₄ paraffins. This can include, but is not limited to, C₃ and/or C₄ paraffins from refinery sources; C₃ and/or C₄ paraffins from chemical production plant sources; and C₃ and/or C₄ paraffins derived from natural gas and/or crude oil stream from a production site, either prior to or after pipeline transport.

The feed including the C₃ and/or C₄ paraffins can also include other components. In various aspects, the feed can include 25 vol % or more of C₃-C₄ paraffins, or 50 vol % or more, or 70 vol % or more such as up to a feed composed substantially of C₃-C₄ paraffins. A feed composed substantially of C₃-C₄ paraffins can have a content of other components of 5.0 vol % or less, or 1.0 vol % or less, or 0.1 vol % or less. For the other components present in the feed, in some aspects the feed can correspond to a feed include C₃₊ paraffins, such as C₃-C₉ paraffins, or C₃-C₆ paraffins. Additionally or alternately, the feed can methane (C₁ paraffin) and/or ethane (C₂ paraffin). Under the conditions in the side riser, methane and ethane remain substantially unreacted, and therefore act as diluents in the feed. Other potential diluents in the feed can include CO₂ and N₂. In some aspects, the feed can include C₄ paraffins, such as C₄₊ paraffins. In such aspects, the feed can include 25 vol % or more of C₄ paraffins, or 50 vol % or more, or 70 vol % or more such as up to a feed composed substantially of C₄ paraffins.

Prior to introducing the feed into the side riser, the feed can be pressurized and/or heated to assist with achieving the desired reaction conditions within the side riser for conversion of paraffins to olefins. For example, the feed can be pressurized to a pressure of 30 psia to 90 psia (˜200 kPa-a to ˜620 kPa-a), or 40 psia to 60 psia (˜270 kPa-a to ˜420 kPa-a). This pressure corresponds to a desired inlet pressure for performing the paraffin to olefin conversion reaction. The feed can be heated to a temperature of 750° F. (˜400° C.) to 950° F. (˜510° C.), or 800° F. (˜425° C.) to 900° F. (˜485° C.).

The heat transfer particles for the side riser can correspond to particles from the separate process associated with the side riser. For example, in an aspect where a side riser is integrated with a fluidized coking process, the heat transfer particles can correspond to coke particles and/or any other particles used in the fluidized coking process for transfer of heat from the heater/burner/gasifier to the coking reactor. In another aspect where a side riser is integrated with a fluid catalytic cracking (FCC) process, the heat transfer particles can correspond to the FCC catalyst particles. The heat transfer particles can be received from the regenerator, gasifier, or other burner for the separate FCC process. After providing heat for the conversion of light paraffins to olefins, the heat transfer particles can be returned to the separate process in any convenient manner. In some aspects, the heat transfer particles can be returned to a primary reactor for the separate process, such as the coking reactor for a fluidized coking process or the FCC reactor for a FCC process. In other aspects, the heat transfer particles can be returned to a secondary vessel, such as a regenerator for an FCC process or a heater in a fluidized coking configuration that includes a coking reactor, a heater, and a gasifier.

After exiting a heater, regenerator, gasifier, or burner, the heat transfer particles can be at an elevated temperature relative to the feed. The heat transfer particles can provide the additional heat that is needed for the endothermic conversion of alkanes to olefins. After exiting the regenerator or burner, the heat transfer particles can be at a typical temperature for particles exiting from a regenerator, gasifier, or other burner, such as a temperature of 550° C. to 1000° C., or 550° C. to 650° C., or 650° C. to 850° C., or 650° C. to 1000° C., or 850° C. to 1000° C. After combination of the heat transfer particles and the feed, the resulting effluent in the side riser can reach a temperature of roughly 1100° F. (˜590° C.) to 1350° F. (˜735° C.).

The flow rates of the heat transfer particles and the feed in the side riser can be selected to provide both a desired residence time (i.e. contact time) between the heat transfer particles and the feed, as well as providing a desired weight ratio of heat transfer particles to feed. One option for characterizing the weight ratio is based on the weight of catalyst to the weight of hydrocarbons. This can be referred to as a catalyst-to-oil ratio. In various aspects, the catalyst-to-oil weight ratio can be 20 or more, or 40 or more, or 50 or more, such as up to 120 or possibly still higher. Additionally or alternately, the residence time of the feed in the side riser can be from 0.5 seconds to 10 seconds, or 0.8 seconds to 8.0 seconds, or 1.0 seconds to 8.0 seconds.

The conditions in the side riser can be suitable for conversion of 10 wt % to 40 wt % of the C₃ paraffins in the feed to olefins, or 20 wt % to 40 wt %, or 10 wt % to 30 wt %. Additionally or alternately, the conditions in the side riser can be suitable for conversion of 20 wt % to 70 wt % of the C₄₊ paraffins in the feed to olefins, or 20 wt % to 50 wt %, or 30 wt % to 60 wt %.

During the conversion of paraffins to olefins, coke can accumulate on the heat transfer particles. However, due to the relatively low WHSV for the feed, and the fact that only the initial conversion of alkanes to olefins is being performed, the amount of coke added to the heat transfer particles can typically be less than the amount of coke that would be added to the particles during a single pass through the reactor for the separate process (i.e., the associated FCC reactor or fluidized coking reactor). Additionally, the flow rate of heat transfer particles from the side riser to the reactor for the separate process can correspond to a minor portion of the flow rate of particles into the separate process. In various aspects, the weight percent of heat transfer particles that pass through the side riser versus the total amount of heat transfer particles introduced into the reactor for the separate process can be 25 wt % or less, or 20 wt % or less, or 15 wt % or less, or 10 wt % or less, such as down to 1.0 wt % or possibly still lower. As a result, any coke added to the heat transfer particles can be managed by the regenerator/gasifier/burner for the separate process.

After the conversion process in the side riser, the resulting effluent can be passed into a separation stage for separation of the heat transfer particles from the olefin-containing gas phase effluent. Cyclone separators are a suitable type of separator for the separation stage. Any convenient configuration of cyclone separators can be used, such as a series of cyclones in series and/or in parallel. Optionally, the separation stage can further include a catalyst wash in order to remove entrained particles fines from the gas phase effluent. Examples of a suitable catalyst wash can correspond to a water wash tower or a vacuum gas oil (VGO) wash tower. The gas phase effluent can then be passed to a second reaction stage that includes a conversion catalyst for conversion of olefins to aromatics.

Reaction System Configuration

FIG. 1 shows an example of a reactor configuration for conversion of paraffins into liquid products. In FIG. 1, a feed 110 that includes C₃ and/or C₄ paraffins is introduced into a processing block for a process such as fluid catalytic cracking (FCC) or fluidized coking. The feed 110 can be heated in a heater 115 (or optionally a heat exchanger) prior to passing the feed into a side riser 190. Optionally, a vaporizer 112 can be included to vaporize the feed 110 prior to heating in heater 115. Side riser 190 also receives a flow of heat transfer particles from a separate processing system that includes a reactor 170 and a burner 180. Reactor 170 can correspond to, for example, a fluidized coking reactor or a FCC reactor. Burner 180 can correspond to, for example, a gasifier or burner for a fluidized coker, or a regenerator associated with an FCC system. It is noted that in some aspects, heat transfer particles from a burner 180 can pass through another vessel prior to entering the side riser 190. For example, in a fluidized coking system including a coking reactor, a heater, and a gasifier, it may be desirable to withdraw heat transfer particles from the heater rather than the gasifier. The heat transfer particles can be passed from burner 180 to side riser 190 via conduit 185. In FIG. 1, a separate conduit 181 is used to return heat transfer particles from burner 180 to reactor 170. Alternatively, conduit 185 could correspond to a side stream from conduit 181. After use in the reactor 170, heat transfer particles (such as coke particles or FCC catalyst particles) can be passed to the burner 180 via conduit 171.

The feed is contacted with the heat transfer particles in side riser 190 under effective conditions for converting a portion of the C₃-C₄ paraffins in the feed to olefins. The resulting effluent 199 from the side riser 190 is then passed into a separation stage, such as one or more cyclones 192 and a (optional) particle wash 194 to produce an olefin-containing (gas phase) effluent 195. Heat transfer particles separated from effluent 199 in the one or more cyclones 192 can be returned to the separate process. In the configuration shown in FIG. 1, the heat transfer particles are returned to the reactor 170 via conduit 172. In other aspects, the heat transfer particles can be returned to any convenient location in the separate process.

The gas phase effluent 195 is then optionally heated (such as in heater 115) prior to being passed into an aromatic formation process. Although FIG. 1 shows the same heater 115 being used for heating the gas phase effluent 195, in other aspects separate heaters and/or heat exchangers can be used for modifying the temperature of gas phase effluent 195 and feed 110. In various aspects, any convenient type of reactor configuration can be used for the aromatic formation process.

The example shown in FIG. 1 includes an aromatic formation process based on a two-train fixed-bed reactor configuration including 2 stages in each train. In other aspects, any convenient number of stages can be included in an aromatic formation process train. In the example shown in FIG. 1, the fixed-bed reactors correspond to radial flow reactors containing a conversion catalyst with aromatization activity, but any convenient type of reactor configuration can be used, such as axial flow reactors. During operation, one reactor train can be operated for conversion of the gas phase effluent (containing paraffins and olefins) to liquids while the other reactor train is undergoing a regeneration cycle. The flows to the respective reactor trains can be controlled using valves that allow paraffin-containing feed or air into the trains as appropriate to the portion of the cycle the reactor train is currently performing.

After any optional heating (such as in heater 115), the gas phase effluent 195 is converted in a series of reactors 140 and 142. In other aspects, additional reactors can be included in the series of reactors. As described below, using a plurality of reactors in sequence can allow for improved control over the reaction temperature and pressure. The control over the reaction temperature can be facilitated in part by the use of intermediate heater 135 located between reactors 140 and 142. In the more general case, any convenient number of heaters and/or heat exchangers can be used to provide temperature control between a plurality of stages in the series of reactors. After performing conversion, the resulting converted effluent 145 is passed to a product separation stage 120. This can allow for separation of the effluent into, for example, liquid products 125, and one or more light gas products 129. Light gas products 129 can potentially include CO₂, C⁴⁻ hydrocarbons, and Hz. Examples of uses for light gas products 129 can include, but are not limited to, using the light gas products as fuel and/or passing the light gas products 129 into a separator to recover hydrogen. Such recovered hydrogen can optionally be used elsewhere in a refinery setting.

It is noted that reactors 150 and 152 similarly correspond to a series of reactors, with intermediate heater 135 between reactors 150 and 152. The difference between reactors 140 and 142 versus reactors 150 and 152 is the portion of the cycle being performed in each reactor train. In the example shown in FIG. 1, the reactor train corresponding to reactors 140 and 142 is in the aromatic formation portion of the cycle, while the reactor train corresponding to reactors 150 and 152 is in the regeneration portion of the cycle. During regeneration, air 161 (or another oxygen-containing gas) is passed via regeneration stage 160 into regeneration input flow line 165. The regeneration stage 160 controls the oxygen concentration in the regeneration input flow line 165. Initially, the regeneration input flow line can include a lower O₂ concentration. The O₂ concentration can be gradually increased up to a final desired value, such as up to roughly 7.0 vol % O₂. The regeneration input flow 165 is heated to allow for regeneration in reactors 150 and 152 at elevated temperature in the presence of oxygen, so that coke can be burned off of the conversion catalyst. Gradually ramping up the O₂ concentration can allow for control of the exotherm that is generated by burning of coke during regeneration. The flue gas generated during regeneration can also be passed through the regeneration block 160. Regeneration products corresponding to water and CO_(x) can be removed from the flue gas, respectively, by water exhaust 169 and vent 167.

In the reactor trains shown in FIG. 1, reactor 140 is in indirect fluid communication with reactor 142, as the effluent from reactor 140 passes through an intervening process heater 135 prior to entering reactor 142. Reactor 140 is considered in direct fluid communication with process heater 135. Reactors 150 and 152 have a similar relationship to the fluid communication relationships between reactors 140 and 142.

FIG. 2 shows another example of a reactor configuration for an aromatic formation process corresponding to a moving bed configuration. A moving bed configuration can be beneficial, relative to configurations such as fixed bed configurations, based on a more stable yield, less variation in conversion, and/or a lower pressure drop. Additionally, the moving bed reactor configuration generally involves a smaller number of distinct reactor vessels and does not require a swing of reactors between production and regeneration. In FIG. 2, gas phase effluent 195 (optionally after heating in a heater, such as heater 115) is passed into the first stage 240 of a moving bed reactor. The reactor in FIG. 2 is shown as including two reactor stages, but any convenient number of stages can be used. The moving bed reactor can either be a radial flow reactor or an axial flow reactor. For the aromatic formation process, a radial flow reactor can be beneficial by providing a reduced pressure drop across the reactor. The reactor stages of the moving bed reactor can be vertically stacked or horizontally stacked. In a vertically stacked configuration, as shown in FIG. 2, the conversion catalyst with aromatization activity can move downward from an earlier stage to a later stage through gravity. The flow of conversion catalyst can be controlled with a valve at the bottom of the final stage. Thus, in the example shown in FIG. 2, conversion catalyst can pass from first stage 240 to second stage 250 via gravity. This type of configuration can be beneficial for reducing catalyst attrition.

In a moving bed reactor, as the conversion catalyst with aromatization activity moves through the reactor and eventually exits from the final stage, the conversion catalyst is continuously replaced with additional catalyst entering at the beginning of the first stage. The replacement catalyst can be provided (at least in part) as regenerated conversion catalyst that is delivered from regenerator 260 to first stage 240 via conduit 265. After exiting from the final stage 250, conversion catalyst can be passed 255 into the regenerator 260. The overall turnaround of the conversion catalyst can be between 1-3 days, or possibly more, depending on the severity of the aromatic formation conditions. The severity of the aromatic formation conditions can determine the rate of coking and catalyst deactivation. It is noted that the rate of coking (and corresponding deactivation due to coking) is reduced or minimized due to the conversion of at least a portion of the paraffins to olefins in the prior side riser 190.

During operation of the aromatic formation process, gas phase products from the first stage 240 can be separated at the bottom of the first stage using some sort of stripping gas. The separated gas phase products can then be passed into interstage furnace 235 before being reintroduced into the top of second stage 250. The product 245 from the final stage can then be cooled and sent to product separation block 120.

Aromatic Formation Processing Conditions

In various aspects, an aromatic formation process can be performed by exposing a mixed feed containing light olefins and light paraffins (such as the effluent from the side riser) to a conversion catalyst with aromatization activity in the presence of hydrogen at suitable conversion conditions. It has been determined that reaction conditions including a temperature of about 500° C. to about 650° C. (or about 550° C. to 650° C., or about 525° C. to about 625° C., or about 550° C. to about 600° C.), a total pressure of about 100 kPa-a to about 525 kPa-a (or about 200 kPa-a to about 525 kPa-a, or about 200 kPa-a to about 450 kPag, or about 300 kPa-a to about 450 kPa-a, or about 200 kPa-a to about 350 kPa-a), and a weight hour space velocity (WHSV) of about 0.1 hr⁻¹ to about 4.0 hr⁻, or about 0.5 hr⁻¹ to about 2.0 hr⁻¹, can be beneficial for performing paraffin conversion to liquid (aromatic) products. Additionally, maintaining a partial pressure of hydrogen within the reactor can be beneficial for reducing or minimizing the formation of coke on the conversion catalyst. Additionally or alternately, if the reaction temperature drops below 450° C., or alternatively below 475° C., the reaction rate of the desired paraffin-to-aromatics conversion reaction can drop significantly, while increasing the temperature above about 650° C. can lead to increases in the catalyst coking rate.

In order to maintain a desired level of control over the temperature during the conversion reaction, the conversion catalyst with aromatization activity can be distributed across multiple reactors in a reactor train. For example, in a configuration that includes two reactors, a first reactor can include 20 wt % to 40 wt % of the conversion catalyst while the second reactor can include the remaining balance of the catalyst. During operation, the olefin-containing effluent from the side riser can be initially heated to a desired temperature, such as 500° C. or more, or 550° C. or more, or 600° C. or more, such as up to 650° C. The feed can then be passed into a first reactor. Due to the highly endothermic nature of aromatic formation from a paraffin feed, the temperature drop across the first reactor in a reactor train can potentially be difficult to manage. However, by performing initial conversion of a portion of paraffins to olefins in the side riser in the presence of heat transfer particles, concerns about excessive temperature drop in the first reactor can be reduced or minimized. After the first reactor, an intermediate heater can be used to return the temperature of the partially converted feed to a desired temperature, such as at least 550° C. or at least 575° C., such as up to about 600° C. or up to about 625° C. The second reactor can then include another portion of the conversion catalyst that is exposed to the feed to complete a desired amount of conversion on the components in the feed. The effluent from the final reactor can then be passed into a separation stage, where liquid aromatic products (C₆₊) and/or other liquid products can be separated from light gases (C₁-C₃ and H₂) in the effluent.

After processing for a period of time, the conversion catalyst (with aromatization activity) in the reactor train can accumulate coke, which can result in a loss in catalyst activity. To restore at least a portion of the catalyst activity the conversion catalyst can be regenerated. In a configuration such as FIG. 1, the reactor train can be switched to a regeneration mode to remove coke from the conversion catalyst. After optionally purging the catalyst train of any remaining hydrocarbons, diluted air having an oxygen content of about 0.5 vol % to about 7 vol % (or another oxygen-containing gas) can be heated to a temperature of between 325° C. to 600° C. and introduced into the reaction train. The heated oxidizing environment can allow for combustion of coke to form carbon oxides and some water, which can then be vented and/or recovered from the reaction train.

In a configuration such as FIG. 2, where the conversion catalyst is continuously removed from the reactor, regeneration of the conversion catalyst can be performed in a separate regeneration vessel. The conditions during regeneration can otherwise be similar. After regeneration, the regenerated conversion catalyst (and optionally a make-up portion to replace catalyst losses) can be returned to the top of the moving bed in the initial reactor stage.

The conversion catalyst (with aromatization activity) can include at least one metal component on an inorganic support, such as amorphous silica, or alumina. The inorganic support may be a porous material such as a micro-porous crystalline material or a meso-porous material. Additionally, suitable molecular sieves may be utilized in the present catalyst and may include at least one medium pore molecular sieve having a Constraint Index of 2-12 (as defined in U.S. Pat. No. 4,016,218). Examples of such medium pore molecular sieves include ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, MCM-22 and MCM-49 and mixtures and intermediates thereof. ZSM-5 is described in detail in U.S. Pat. No. 3,702,886 and Re. 29,948. ZSM-11 is described in detail in U.S. Pat. No. 3,709,979. A ZSM-5/ZSM-11 intermediate structure is described in U.S. Pat. No. 4,229,424. ZSM-12 is described in U.S. Pat. No. 3,832,449. ZSM-22 is described in U.S. Pat. No. 4,556,477. ZSM-23 is described in U.S. Pat. No. 4,076,842. ZSM-35 is described in U.S. Pat. No. 4,016,245. ZSM-48 is more particularly described in U.S. Pat. No. 4,234,231

The metal component of the conversion catalyst may be present in an amount of at least 0.1 wt. %, such as from 0.1 to 5 wt. %, of the overall catalyst. The metal component may include one or more neutral metals selected from Groups 3 to 13 of the Periodic Table of the Elements, such as Ga, In, Zn, Cu, Re, Mo, W, La, Fe, Ag, Pt, Pd, and/or one or more oxides, sulfides and/or carbides of these metals. The metal component can be provided on the conversion catalyst in any known manner, for example by impregnation or ion exchange of the molecular sieve with a solution of a compound of the relevant metal, followed by conversion of the metal compound to the desired form, namely neutral metal, oxide, sulfide and/or carbide. Part or all of the metal may also be present in a crystalline framework of the molecular sieve.

In a preferred embodiment, a bifunctional conversion catalyst (with aromatization activity) may be selected from the group consisting of Ga and/or In-modified ZSM-5 type zeolites, such as Ga and/or In-impregnated H-ZSM-5, Ga and/or In-exchanged H-ZSM-5, H-gallosilicate of ZSM-5 type structure and H-galloaluminosilicate of ZSM-5 type structure. These zeolites can also be prepared by methods known in the prior art.

For example, the bifunctional conversion catalyst may contain tetrahedral aluminum or gallium, which is present in the zeolite framework or lattice. The bifunctional conversion catalyst may also contain octahedral gallium or indium, which is not present in the zeolite framework, but present in the zeolite channels in close vicinity to the zeolitic protonic acid sites that may be attributed to the presence of tetrahedral aluminum and gallium in the catalyst. The tetrahedral or framework of Al or Ga can be responsible for the acid function of the catalyst and octahedral or non-framework Ga or In may be responsible for the dehydrogenation function of the catalyst. In a preferred embodiment, the bifunctional conversion catalyst may include H-galloaluminosilicate of ZSM-5 type structure having framework (tetrahedral) Si/Al and Si/Ga mole ratios of about 10:1 to 100:1 and 15:1 to 150:1, respectively, and non-framework (octahedral) Ga of about 0.5 to 0 wt. %.

In addition to the molecular sieve and hydrogenation component, the conversion catalyst may be composited with other materials, which may be resistant to the temperatures and other conditions employed in the conversion reaction. Such other materials can include active and inactive materials and synthetic or naturally occurring zeolites as well as inorganic materials such as clays and/or oxides such as alumina, silica, silica-alumina, zirconia, titania, magnesia or mixtures of these and other oxides. The latter may be either naturally occurring or in the form of gelatinous precipitates or gels including mixtures of silica and metal oxides. Clays may also be included with the oxide type binders to modify the mechanical properties of the catalyst or to assist in its manufacture. Use of a material in conjunction with the molecular sieve, i.e., combined therewith or present during its synthesis, which itself is catalytically active may change the conversion and/or selectivity of the conversion catalyst.

Additionally, inactive materials may serve as diluents to control the amount of conversion so that products can be obtained without employing other means for controlling the rate of reaction. These materials may be incorporated into naturally-occurring clays, e.g., bentonite and kaolin, to improve the crush strength of the catalyst under commercial operating conditions and function as binders or matrices for the catalyst. The relative proportions of molecular sieve and inorganic oxide matrix may vary, with the sieve content ranging from about 1 to about 90 percent by weight and more usually, particularly, when the composite is prepared in the form of beads, in the range of about 2 to about 80 weight percent of the composite.

In any aspect, the conversion catalyst (with aromatization activity) may comprise least one molecular sieve component and at least one dehydrogenation component. The molecular sieve component may comprise >80 wt. % to <100 wt. %, preferably 100 wt. % of said catalyst. In one or more embodiments, the conversion catalyst can be substantially free of binder, such as an inorganic binder, or matrix material, e.g., contains ≤1 wt. % of binder or matrix material, such as, for example, ≤0.1 wt. %, based on the weight of the catalyst. In other embodiments, the molecular sieve is present in the catalyst in an amount of >80 wt. %, ≥90 wt. %, or ≥95 wt. %, or ≥98 wt. %, or ≥99 wt. %, based on the weight of the catalyst. In some embodiments, the molecular sieve is an aluminosilicate and is present in the catalyst in a range of from about 80 wt. %, or 85 wt. %, or 90 wt. %, or 95 wt. % up to about 99.9 wt. %, based on the weight of the catalyst. In a most preferred embodiment, the molecular sieve is an aluminosilicate and present in the catalyst in the amount of about 100 wt. %. The molecular sieve may consist essentially of or even consist of an aluminosilicate.

The crystalline aluminosilicate of the conversion catalysts may have a constraint index of less than 12, preferably, in the range of about 1 to about 12. Typically, the crystalline aluminosilicate is one having a medium pore size and a Constraint Index of less than or equal to about 12. Constraint Index is defined in U.S. Pat. No. 4,016,218. Examples of suitable aluminosilicates include ZSM-5, ZSM-11, ZSM-12, ZSM-21, ZSM-22, ZSM-23, ZSM-35, ZSM-38, ZSM-48 ZSM-50, ZSM-57, and MCM-68, including mixtures and intermediates thereof such as ZSM-5/ZSM-11 admixture. ZSM-5 is described in U.S. Pat. Nos. 3,702,886 and Re. 29,948. ZSM-11 is described in U.S. Pat. No. 3,709,979. A ZSM-5/ZSM-11 intermediate structure is described in U.S. Pat. No. 4,229,424. ZSM-12 is described in U.S. Pat. No. 3,832,449. ZSM-21 is described U.S. Pat. No. 4,082,805. ZSM-22 is described in U.S. Pat. No. 4,556,477. ZSM-23 is described in U.S. Pat. No. 4,076,842. ZSM-35 is described in U.S. Pat. No. 4,016,245. ZSM-38 is described in U.S. Pat. No. 4,046,859. ZSM-48 is described in U.S. Pat. No. 4,234,231. ZSM-50 is described in U.S. Pat. No. 4,640,826. ZSM-57 is described in U.S. Pat. No. 4,873,067. TEA-Mordenite is described in U.S. Pat. Nos. 3,766,093 and 3,894,104. MCM-68 is described in U.S. Pat. No. 6,049,018.

The aluminosilicate's silica-to-alumina (Si:Al₂) atomic ratio may be typically ≥2 molar, e.g., in the range of 10 to 300 molar, or in the range of from 5 to 100 molar. The silica-to-alumina ratio, Si:Al₂, is meant to represent the Si:Al₂ atomic ratio in the rigid anionic framework of the crystalline aluminosilicate. In other words, aluminum in (i) any matrix or binder or (ii) in cationic or other form within the crystalline aluminosilicate's channels is excluded from the Si:Al₂ atomic ratio. Aluminosilicates having a higher silica-to-alumina ratio can be utilized when a lower catalyst acidity is desired, e.g., in the range of from 44 to 100 molar, such as from 50 to 80 molar, or 55 to 75 molar.

In one or more embodiments, the crystalline aluminosilicate has a constraint index in the range of about 1 to 12 and is selected from the group consisting of a MCM-22 family material,

ZSM-5, ZSM-11, ZSM-12, ZSM-21, ZSM-22, ZSM-23, ZSM-35, ZSM-38, ZSM-48, ZSM-50, ZSM-57, MCM-68 and mixtures of two or more thereof. Preferably, the aluminosilicate is ZSM-11 or H-ZSM-11 (the acidic form of ZSM-11), and more preferably, the aluminosilicate is ZSM-5 or H-ZSM-5 (the acidic form of ZSM-5).

In certain aspects, the molecular sieve has a relatively small crystal size, e.g., small crystal ZSM-5, meaning ZSM-5 has a crystal size ≤0.05 μm, such as in the range of 0.02 μm to 0.05 μm. Small crystal ZSM-5 and the method for determining molecular sieve crystal sizes are disclosed in U.S. Pat. No. 6,670,517, which is incorporated by reference herein in its entirety.

In other aspects, the crystalline aluminosilicate comprises at least one molecular sieve of the MCM-22 family, e.g., MCM-22 alone or in combination with other aluminosilicates, specified above, or other MCM-22 family materials. Materials of the MCM-22 family include MCM-22 (described in U.S. Pat. No. 4,954,325), PSH-3 (described in U.S. Pat. No. 4,439,409), SSZ-25 (described in U.S. Pat. No. 4,826,667), ERB-1 (described in European Patent No. 0293032), ITQ-1 (described in U.S. Pat. No. 6,077,498), and ITQ-2 (described in International Patent Publication No. WO97/17290), MCM-36 (described in U.S. Pat. No. 5,250,277), MCM-49 (described in U.S. Pat. No. 5,236,575), MCM-56 (described in U.S. Pat. No. 5,362,697) and mixtures of two or more thereof. Related aluminosilicates to be included in the MCM-22 family are UZM-8 (described in U.S. Pat. No. 6,756,030) and UZM-8HS (described in U.S. Pat. No. 7,713,513), both of which are also suitable for use as the molecular sieve component.

In one or more embodiments, the molecular sieve may be one that is in hydrogen form, e.g., one that has been synthesized in the alkali metal form, but is then converted from the alkali to the hydrogen form and has hydrogen ions, e.g., acidic.

The dehydrogenation component comprises at least one element from Group 5 to 15 of the Periodic Chart. In one or more embodiments, the element is a metal. Preferably, the catalyst comprise at least one first metal. The first metal is selected from the group consisting of and includes one or more of zinc, gallium, copper, silver, tin, iron, cobalt, nickel, gold, manganese, chromium, molybdenum, tungsten, and mixtures of two or more thereof. Preferably, the first metal is zinc or gallium.

In one or more embodiments, the dehydrogenation component of the conversion catalysts of this invention further comprises at least one second metal in addition to the first metal. The second metal is different from the first metal. The second metal is selected from the group consisting of and includes one or more of phosphorus, platinum, palladium, lanthanum rhenium, and mixtures of two or more thereof. Preferably, the second metal is phosphorous.

The conversion catalyst contains at least about 0.005 wt. % of the first metal, or in the range from about 0.005 wt. % to about 4.0 wt. % of said first metal, or from about 0.01 wt. % to about 3.0 wt. % of said first metal, based on the weight of said catalyst. When the second metal is present, the conversion catalyst contains in the range from 0 wt. % to about 5.0 wt. % of the second metal, or from about 0.005 wt. % to about 4.0 wt. % of said second metal, or from about 0.01 wt. % to about 3.0 wt. % of said second metal, based on the weight of the catalyst.

Not being bound by any theory, it is believed that the higher catalyst activity provided by the increased molecular sieve (preferably, aluminosilicate) content is modulated by the presence of the first metal and optionally, the second metal, leading to increased cycle length when the catalyst is used in a process for conversion of a light paraffinic hydrocarbon feedstock in a fixed bed reactor.

An inorganic binder or matrix material may be used, e.g., to make the conversion catalyst more resistant to the temperatures and other conditions employed in the process for conversion. The amount of such inorganic binder or matrix material is set forth above.

The inorganic binder or matrix material may include clays and/or inorganic oxides. Such inorganic binders include alumina, silica, silica-alumina, titanic, zirconia, magnesia, tungsten oxide, ceria, niobia, and mixtures of two or more thereof The matrix component may include naturally occurring materials and/or materials in the form of gelatinous precipitates or gels including mixtures of silica and metal oxides. Clays may also be included with the oxide type binders to modify the mechanical properties of the catalyst or to assist in its manufacture. Alternatively, or in addition, the inorganic binder or matrix material may include one or more substantially inactive materials. Inactive materials suitably serve as diluents to control the amount of conversion so that products may be obtained economically and orderly without employing other means for controlling the rate of reaction. These materials may be incorporated into naturally occurring clays, e.g., bentonite and kaolin, to improve thermal and strength properties (e.g., crush strength) of the catalyst under catalytic conversion conditions. The binder or matrix material may include active materials, such as synthetic or naturally occurring aluminosilicate.

The conversion catalyst, preferably a conversion catalyst having a molecular sieve component comprising >80 wt. % to <100 wt. % of said catalyst and a dehydrogenation component comprised of a first metal and an optional second metal, provides ≤a 5% reduction in yield of said aromatic hydrocarbon after a time-on-stream of ≤50 hours.

In one or more embodiments, the conversion catalyst produces ≤a 5% reduction, or a 7% reduction, preferably ≤a 10% reduction, in the yield of the aromatic hydrocarbon after the time-on-stream of the catalyst is at least 20 hours, or at least 50 hours, preferably at least 100 hours under conversion conditions of a temperature in the range of about 450° C. to about 750° C., a pressure in the range of from about 35 kPa to about 1480 kPa and a WHSV from about 0.1 to about 20 hr⁻¹.

In a preferred embodiment, the conversion catalyst of this invention is a catalyst for producing aromatic hydrocarbon comprising: (a) from >80 wt. % to <100 wt. % of a crystalline aluminosilicate which comprises ZSM-5 or ZSM-11, based on the weight of said catalyst; (b) about 0.005 wt. % to about 5.0 wt. % of a first metal, based on the weight of said catalyst, and (c) 0 wt. % to about 5.0 wt. % of a second metal, based on the weight of said catalyst, wherein said catalyst has ≤a 5% reduction in yield of aromatic hydrocarbon after an average time-on-stream in said reaction zone of at least 50 hours in a process for producing aromatic hydrocarbon which includes converting ≥1 wt. % of a C₂₊ non-aromatic hydrocarbon in the presence of said catalyst to a product comprising ≥1 wt. % of aromatic hydrocarbon under conversion conditions that include a temperature in the range of about 450° C. to about 750° C., a pressure in the range of from about 35 kPa to about 1480 kPa and a WHSV from about 0.1 to about 20 hr⁻¹. In one or more embodiments, the first metal, optional second metal and inorganic binder are as set forth above. The conversion catalysts of this invention are made by any one of the methods disclosed hereinafter.

A method of making one or more conversion catalysts (with aromatization activity) for use in the process may comprise the first step of providing a molecular sieve, preferably a crystalline aluminosilicate, which has a constraint index of less than or equal to about 12, preferably in the range of about 1 to about 12, more preferably, a crystalline aluminosilicate comprising ZSM-5 or ZSM-11. In a contacting step, the molecular sieve is contacted with a source of a first metal and optionally a source of a second metal under conditions sufficient to deposit said first metal and said optional second metal on the molecular sieve and to form a metal-containing molecular sieve.

If both metals are used, the first metal is different from the second metal. The first metal is selected from the group consisting of zinc, gallium, copper, silver, tin, iron, cobalt, nickel, gold, manganese, chromium, molybdenum, indium, tungsten, and mixtures of two or more thereof. The second metal is selected from the group consisting of phosphorus, platinum, palladium, lanthanum rhenium, and mixtures of two or more thereof.

The first metal and the optional second metal may be deposited on the crystalline aluminosilicate, by any suitable method, e.g., by impregnating the first metal and optionally the second metal onto the external surface of the molecular sieve, preferably the crystalline aluminosilicate. The first metal and optionally the second metal may be dissolved in a liquid carrier, for example an aqueous or organic carrier, mixed with the catalyst, and then dried by evaporation or vacuum distillation. This method may be termed “impregnation”. Other conventional methods may be utilized to deposit the first metal and the optional second metal onto the molecular sieve, preferably the crystalline aluminosilicate, such as for example, by the incipient wetness method, and the invention is not limited to any one specific method. Non-limiting examples of the conditions effective to deposit the first metal and optionally the second metal on the molecular sieve, preferably the crystalline aluminosilicate, are set forth in Example 3.

When the first metal is zinc, non-limiting suitable sources of zinc are selected from the group consisting of zinc nitrate, zinc titanate, zinc silicate, zinc borate, zinc fluorosilicate, zinc fluorotitanate, zinc molybdate, zinc chromate, zinc tungstate, zinc zirconate, zinc chromite, zinc aluminate, zinc phosphate, zinc acetate dihydrate, diethyl zinc, zinc 2-ethylhexanoate, and mixtures of two or more thereof.

When the second metal is lanthanum, non-limiting suitable sources of lanthanum include a lanthanum salt, a lanthanum nitrate, or a mixture thereof.

Integration of Side Riser with Fluidized Coking

In some aspects, the side riser for converting paraffins to olefins in the presence of heat transfer particles can be integrated with a fluidized coking system. In such aspects, the heat transfer particles for the side riser can correspond to coke particles and/or other particles used for heat transfer in the fluidized coking system. In such aspects, a burner or gasifier for the fluidized coking system can correspond to the burner that provides heat transfer particles to the side riser. For example, in a configuration such as FIG. 1 or FIG. 2, the reactor 170 can correspond to the fluidized coking reactor while the burner 180 can correspond to the burner or gasifier for the fluidized coking system.

In this description, the term “Flexicoking” (trademark of ExxonMobil Research and Engineering Company) is used to designate a fluid coking process in which heavy petroleum feeds are subjected to thermal cracking in a fluidized bed of heated solid particles to produce hydrocarbons of lower molecular weight and boiling point along with coke as a by-product which is deposited on the solid particles in the fluidized bed. The resulting coke can then converted to a fuel gas by contact at elevated temperature with steam and an oxygen-containing gas in a gasification reactor (gasifier). This type of configuration can more generally be referred to as an integration of fluidized bed coking with gasification. FIGS. 3 and 4 provide examples of fluidized coking reactors that include a gasifier.

FIG. 3 shows an example of a Flexicoker unit (i.e., a system including a gasifier that is thermally integrated with a fluidized bed coker) with three reaction vessels: reactor, heater and gasifier. Additionally, FIG. 3 shows integration with a side riser.

In FIG. 3, the Flexicoking unit comprises reactor section 10 with the coking zone and its associated stripping and scrubbing sections (not separately indicated), heater section 11 and gasifier section 12. The relationship of the coking zone, scrubbing zone and stripping zone in the reactor section is shown, for example, in U.S. Pat. No. 5,472,596, to which reference is made for a description of the Flexicoking unit and its reactor section. A heavy oil feed is introduced into the unit by line 13 and cracked hydrocarbon product withdrawn through line 14. Fluidizing and stripping steam is supplied by line 15. Cold coke is taken out from the stripping section at the base of reactor 10 by means of line 16 and passed to heater 11. The term “cold” as applied to the temperature of the withdrawn coke is, of course, decidedly relative since it is well above ambient at the operating temperature of the stripping section. Hot coke is circulated from heater 11 to reactor 10 through line 17. Coke from heater 11 is transferred to gasifier 12 through line 21 and hot, partly gasified particles of coke are circulated from the gasifier back to the heater through line 22. The excess coke is withdrawn from the heater 11 by way of line 23. In conventional configurations, gasifier 12 is provided with its supply of steam and air by line 24 and hot fuel gas is taken from the gasifier to the heater though line 25. In some alternative aspects, instead of supplying air via a line 24 to the gasifier 12, a stream of oxygen with 95 vol % purity or more can be provided, such as an oxygen stream from an air separation unit. In such aspects, in addition to supplying a stream of oxygen, a stream of an additional diluent gas can be supplied by line 31. The additional diluent gas can correspond to, for example, CO2 separated from the fuel gas generated during the gasification. The fuel gas is taken out from the unit through line 26 on the heater; coke fines are removed from the fuel gas in heater cyclone system 27 comprising serially connected primary and secondary cyclones with diplegs which return the separated fines to the fluid bed in the heater. The fuel gas from line 26 can then undergo further processing. For example, in some aspects, the fuel gas from line 26 can be passed into a separation stage for separation of CO₂ (and/or H₂S). This can result in a stream with an increased concentration of synthesis gas, which can then be passed into a conversion stage for conversion of synthesis gas to methanol.

It is noted that in some optional aspects, heater cyclone system 27 can be located in a separate vessel (not shown) rather than in heater 11. In such aspects, line 26 can withdraw the fuel gas from the separate vessel, and the line 23 for purging excess coke can correspond to a line transporting coke fines away from the separate vessel. These coke fines and/or other partially gasified coke particles that are vented from the heater (or the gasifier) can have an increased content of metals relative to the feedstock. For example, the weight percentage of metals in the coke particles vented from the system (relative to the weight of the vented particles) can be greater than the weight percent of metals in the feedstock (relative to the weight of the feedstock). In other words, the metals from the feedstock are concentrated in the vented coke particles. Since the gasifier conditions do not create slag, the vented coke particles correspond to the mechanism for removal of metals from the coker/gasifier environment. In some aspects, the metals can correspond to a combination of nickel, vanadium, and/or iron. Additionally or alternately, the gasifier conditions can cause substantially no deposition of metal oxides on the interior walls of the gasifier, such as deposition of less than 0.1 wt % of the metals present in the feedstock introduced into the coker/gasifier system, or less than 0.01 wt %.

In configurations such as FIG. 3, the system elements shown in the figure can be characterized based on fluid communication between the elements. For example, reactor section 10 is in direct fluid communication with heater 11. Reactor section 10 is also in indirect fluid communication with gasifier 12 via heater 11.

The configuration shown in FIG. 3 also includes a side riser 91. Side riser 91 in FIG. 3 can correspond, for example, to side riser 190 in FIG. 1 or FIG. 2. In a configuration such as FIG. 1 or FIG. 2, the reactor 170 can correspond to a reactor section 10 from FIG. 3, while the burner 180 can correspond to gasifier 12.

It is noted that a fluidized coking configuration such as the configuration in FIG. 3 could result in some changes to the types of configurations shown in FIG. 1 or FIG. 2. For example, in some aspects, heat transfer particles can be passed from the fluidized coking system to the side riser by passing particles from gasifier 12 to side riser 91 via conduit 52. Additionally or alternately, in some aspects a portion of the heat transfer particles passed into the side riser can be transferred from heater 11 to side riser 91 via conduit 62. Further additionally or alternately, depending on the aspect, the cooled heat transfer particles exiting from side riser 91 can be returned to the reactor 10 (via conduit 77), to the heater 11 (via conduit 67), to the gasifier 12 (via conduit 57), or portions of the catalyst can be returned to any convenient combination of reactor 10, heater 11, and/or gasifier 12.

As an alternative, integration of a fluidized bed coker with a gasifier can also be accomplished without the use of an intermediate heater. In such alternative aspects, the cold coke from the reactor can be transferred directly to the gasifier. This transfer, in almost all cases, will be unequivocally direct with one end of the tubular transfer line connected to the coke outlet of the reactor and its other end connected to the coke inlet of the gasifier with no intervening reaction vessel, i.e. heater. The presence of devices other than the heater is not however to be excluded, e.g. inlets for lift gas etc. Similarly, while the hot, partly gasified coke particles from the gasifier are returned directly from the gasifier to the reactor this signifies only that there is to be no intervening heater as in the conventional three-vessel Flexicoker but that other devices may be present between the gasifier and the reactor, e.g. gas lift inlets and outlets.

FIG. 4 shows an example of integration of a fluidized bed coker with a gasifier but without a separate heater vessel. In the configuration shown in FIG. 4, the cyclones for separating fuel gas from catalyst fines are located in a separate vessel. In other aspects, the cyclones can be included in gasifier vessel 41.

In the configuration shown in FIG. 4, the configuration includes a reactor 40, a main gasifier vessel 41 and a separator 42. The heavy oil feed is introduced into reactor 40 through line 43 and fluidizing/stripping gas through line 44; cracked hydrocarbon products are taken out through line 45. Cold, stripped coke is routed directly from reactor 40 to gasifier 41 by way of line 46 and hot coke returned to the reactor in line 47. Steam and oxygen are supplied through line 48. The flow of gas containing coke fines is routed to separator vessel 42 through line 49 which is connected to a gas outlet of the main gasifier vessel 41. The fines are separated from the gas flow in cyclone system 50 comprising serially connected primary and secondary cyclones with diplegs which return the separated fines to the separator vessel. The separated fines are then returned to the main gasifier vessel through return line 51 and the fuel gas product taken out by way of line 52. Coke is purged from the separator through line 53. The fuel gas from line 52 can then undergo further processing for separation of CO₂ (and/or H₂S) and conversion of synthesis gas to methanol.

FIG. 4 also shows an example of integration of the fluidized coking vessels (reactor 40 and main gasifier vessel 41) with a side riser 90. In FIG. 4, heat transfer particles can be passed from the fluidized coking system to the side riser by passing particles from gasifier 41 to side riser 90 via conduit 82. Additionally or alternately, depending on the aspect, the cooled heat transfer particles exiting from side riser 90 can be returned to the reactor 40 (via conduit 97) or to the main gasifier vessel 41 (via conduit 87), or portions of the catalyst can be returned to any convenient combination of reactor 40 and main gasifier vessel 41.

The coker and gasifier can be operated according to the parameters necessary for the required coking processes. The gasification zone is typically maintained at a high temperature ranging from 850° C. to 1000° C. (˜1560° F. to 1830° F.) and a pressure ranging from 0 kPag to 1000 kPag (˜0 psig to 150 psig), preferably from 200 kPag to 400 kPag (˜30 psig to 60 psig). Steam and an oxygen-containing gas are introduced to provide fluidization and an oxygen source for gasification. In some aspects the oxygen-containing gas can be air. In other aspects, the oxygen-containing gas can have a low nitrogen content, such as oxygen from an air separation unit or another oxygen stream including 95 vol % or more of oxygen, or 98 vol % or more, are passed into the gasifier for reaction with the solid particles comprising coke deposited on them in the coking zone. In aspects where the oxygen-containing gas has a low nitrogen content, a separate diluent stream, such as a recycled CO₂ or H₂S stream derived from the fuel gas produced by the gasifier, can also be passed into the gasifier.

Integration of Side Riser with Fluid Catalytic Cracking

In some aspects, the side riser for converting paraffins to olefins in the presence of heat transfer particles can be integrated with a fluid catalytic cracking (FCC) system. In such aspects, the heat transfer particles for the side riser can correspond to FCC catalyst particles. In such aspects, a regenerator for the FCC system can correspond to the burner that provides heat transfer particles to the side riser. For example, in a configuration such as FIG. 1 or FIG. 2, the reactor 170 can correspond to the FCC reactor while the burner 180 can correspond to the regenerator for the FCC system. Optionally, the FCC system can include a catalyst cooler. In such aspects, the side riser comprises a separate reaction vessel from the catalyst cooler.

Suitable feedstreams for processing in an FCC reactor can include, but are not limited to, feeds boiling in the range of about 430° F. to about 1050° F. (˜221° C. to ˜566° C.), such as gas oils, heavy hydrocarbon oils comprising materials boiling above 1050° F. (˜566° C.); heavy and reduced petroleum crude oil; petroleum atmospheric distillation bottoms; petroleum vacuum distillation bottoms; pitch, asphalt, bitumen, other heavy hydrocarbon residues; tar sand oils; shale oil; liquid products derived from coal liquefaction processes; catalytic slurry oils from an FCC process; and mixtures thereof. The FCC feed may comprise recycled hydrocarbons, such as light or heavy cycle oils.

An example of a suitable reactor for performing an FCC process can be a riser reactor. Within the reactor riser, the FCC feedstream can be contacted with a catalytic cracking catalyst under cracking conditions thereby resulting in spent catalyst particles containing carbon deposited thereon and a lower boiling product stream. The cracking conditions can typically include: temperatures from about 900° F. to about 1060° F. (˜482° C. to ˜571° C.), or about 950° F. to about 1040° F. (˜510° C. to ˜560° C.); hydrocarbon partial pressures from about 10 to 50 psia (˜70-350 kPa-a), or from about 20 to 40 psia (˜140-280 kPa-a); and a catalyst to feed (wt/wt) ratio from about 3 to 8, or about 5 to 6, where the catalyst weight can correspond to total weight of the catalyst composite. Steam may be concurrently introduced with the feed into the reaction zone. The steam may comprise up to about 5 wt % of the feed. In some aspects, the FCC feed residence time in the reaction zone can be less than about 5 seconds, or from about 3 to 5 seconds, or from about 2 to 3 seconds.

In some aspects, the FCC can be operated at low temperature, high conversion conditions. During low temperature operation, the FCC unit can be operated at a temperature from about 850° F. (˜454° C.) to about 950° F. (˜510° C.), or about 850° F. (˜454° C.) to about 920° F. (˜493° C.), or about 850° F. (˜454° C.) to about 900° F. (˜482° C.); hydrocarbon partial pressures from about 10 to 50 psia (˜70-350 kPa-a), or from about 20 to 40 psia (˜140-280 kPa-a); and a catalyst to feed (wt/wt) ratio from about 3 to 8, or about 5 to 6, where the catalyst weight can correspond to total weight of the catalyst composite. Steam may be concurrently introduced with the feed into the reaction zone. The steam may comprise up to about 5 wt % of the feed. The residence time for the input feed can be from about 2 seconds to about 8 seconds, or about 4 seconds to about 8 seconds, or about 4 seconds to about 6 seconds.

Catalysts suitable for use within the FCC reactor herein can be fluid cracking catalysts comprising either a large-pore molecular sieve or a mixture of at least one large-pore molecular sieve catalyst and at least one medium-pore molecular sieve catalyst. Large-pore molecular sieves suitable for use herein can be any molecular sieve catalyst having an average pore diameter greater than ˜0.7 nm which are typically used to catalytically “crack” hydrocarbon feeds. In various aspects, both the large-pore molecular sieves and the medium-pore molecular sieves used herein be selected from those molecular sieves having a crystalline tetrahedral framework oxide component. For example, the crystalline tetrahedral framework oxide component can be selected from the group consisting of zeolites, tectosilicates, tetrahedral aluminophosphates (ALPOs) and tetrahedral silicoaluminophosphates (SAPOs). Preferably, the crystalline framework oxide component of both the large-pore and medium-pore catalyst can be a zeolite. More generally, a molecular sieve can correspond to a crystalline structure having a framework type recognized by the International Zeolite Association. It should be noted that when the cracking catalyst comprises a mixture of at least one large-pore molecular sieve catalyst and at least one medium-pore molecular sieve, the large-pore component can typically be used to catalyze the breakdown of primary products from the catalytic cracking reaction into clean products such as naphtha and distillates for fuels and olefins for chemical feedstocks.

Large pore molecular sieves that are typically used in commercial FCC process units can be suitable for use herein. FCC units used commercially generally employ conventional cracking catalysts which include large-pore zeolites such as USY or REY. Additional large pore molecular sieves that can be employed in accordance with the present invention include both natural and synthetic large pore zeolites. Non-limiting examples of natural large-pore zeolites include gmelinite, chabazite, dachiardite, clinoptilolite, faujasite, heulandite, analcite, levynite, erionite, sodalite, cancrinite, nepheline, lazurite, scolecite, natrolite, offretite, mesolite, mordenite, brewsterite, and ferrierite. Non-limiting examples of synthetic large pore zeolites are zeolites X, Y, A, L. ZK-4, ZK-5, B, E, F, H, J, M, Q, T, W, Z, alpha and beta, omega, REY and USY zeolites. In some aspects, the large pore molecular sieves used herein can be selected from large pore zeolites. In such aspects, suitable large-pore zeolites for use herein can be the faujasites, particularly zeolite Y, USY, and REY.

Medium-pore size molecular sieves that are suitable for use herein include both medium pore zeolites and silicoaluminophosphates (SAPOs). Medium pore zeolites suitable for use in the practice of the present invention are described in “Atlas of Zeolite Structure Types”, eds. W. H. Meier and D. H. Olson, Butterworth-Heineman, Third Edition, 1992, hereby incorporated by reference. The medium-pore size zeolites generally have an average pore diameter less than about 0.7 nm, typically from about 0.5 to about 0.7 nm and includes for example, MFI, MFS, MEL, MTW, EUO, MTT, HEU, FER, and TON structure type zeolites (IUPAC Commission of Zeolite Nomenclature). Non-limiting examples of such medium-pore size zeolites, include ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-34, ZSM-35, ZSM-38, ZSM-48, ZSM-50, silicalite, and silicalite 2. An example of a suitable medium pore zeolite can be ZSM-5, described (for example) in U.S. Pat. Nos. 3,702,886 and 3,770,614. Other suitable zeolites can include ZSM-11, described in U.S. Pat. No. 3,709,979; ZSM-12 in U.S. Pat. No. 3,832,449; ZSM-21 and ZSM-38 in U.S. Pat. No. 3,948,758; ZSM-23 in U.S. Pat. No. 4,076,842; and ZSM-35 in U.S. Pat. No. 4,016,245. As mentioned above SAPOs, such as SAPO-11, SAPO-34, SAPO-41, and SAPO-42, described (for example) in U.S. Pat. No. 4,440,871 can also be used herein. Non-limiting examples of other medium pore molecular sieves that can be used herein include chromosilicates; gallium silicates; iron silicates; aluminum phosphates (ALPO), such as ALPO-11 described in U.S. Pat. No. 4,310,440; titanium aluminosilicates (TASO), such as TASO-45 described in EP-A No. 229,295; boron silicates, described in U.S. Pat. No. 4,254,297; titanium aluminophosphates (TAPO), such as TAPO-11 described in U.S. Pat. No. 4,500,651 and iron aluminosilicates. All of the above patents are incorporated herein by reference.

The medium-pore size zeolites (or other molecular sieves) used herein can include “crystalline admixtures” which are thought to be the result of faults occurring within the crystal or crystalline area during the synthesis of the zeolites. Examples of crystalline admixtures of ZSM-5 and ZSM-11 can be found in U.S. Pat. No. 4,229,424, incorporated herein by reference. The crystalline admixtures are themselves medium-pore size zeolites, in contrast to physical admixtures of zeolites in which distinct crystals of crystallites of different zeolites are physically present in the same catalyst composite or hydrothermal reaction mixtures.

In some aspects, the large-pore zeolite catalysts and/or the medium-pore zeolite catalysts can be present as “self-bound” catalysts, where the catalyst does not include a separate binder. In some aspects, the large-pore and medium-pore catalysts can be present in an inorganic oxide matrix component that binds the catalyst components together so that the catalyst product can be hard enough to survive inter-particle and reactor wall collisions. The inorganic oxide matrix can be made from an inorganic oxide sol or gel which can be dried to “glue” the catalyst components together. Preferably, the inorganic oxide matrix can be comprised of oxides of silicon and aluminum. It can be preferred that separate alumina phases be incorporated into the inorganic oxide matrix. Species of aluminum oxyhydroxides-γ-alumina, boehmite, diaspore, and transitional aluminas such as α-alumina, β-alumina, γ-alumina, δ-alumina, ε-alumina, κ-alumina, and ρ-alumina can be employed. Preferably, the alumina species can be an aluminum trihydroxide such as gibbsite, bayerite, nordstrandite, or doyelite. Additionally or alternately, the matrix material may contain phosphorous or aluminum phosphate. Optionally, the large-pore catalysts and medium-pore catalysts be present in the same or different catalyst particles, in the aforesaid inorganic oxide matrix.

In the FCC reactor, the cracked FCC product can be removed from the fluidized catalyst particles. Preferably this can be done with mechanical separation devices, such as an FCC cyclone. The FCC product can be removed from the reactor via an overhead line, cooled and sent to a fractionator tower for separation into various cracked hydrocarbon product streams. These product streams may include, but are not limited to, a light gas stream (generally comprising C₄ and lighter hydrocarbon materials), a naphtha (gasoline) stream, a distillate (diesel and/or jet fuel) steam, and other various heavier gas oil product streams. The other heavier stream or streams can include a bottoms stream.

In the FCC reactor, after removing most of the cracked FCC product through mechanical means, the majority of, and preferably substantially all of, the spent catalyst particles can be conducted to a stripping zone within the FCC reactor. The stripping zone can typically contain a dense bed (or “dense phase”) of catalyst particles where stripping of volatiles takes place by use of a stripping agent such as steam. There can also be space above the stripping zone with a substantially lower catalyst density which space can be referred to as a “dilute phase”. This dilute phase can be thought of as either a dilute phase of the reactor or stripper in that it will typically be at the bottom of the reactor leading to the stripper.

In some aspects, the majority of, and preferably substantially all of, the stripped catalyst particles are subsequently conducted to a regeneration zone wherein the spent catalyst particles are regenerated by burning coke from the spent catalyst particles in the presence of an oxygen containing gas, preferably air thus producing regenerated catalyst particles. This regeneration step restores catalyst activity and simultaneously heats the catalyst to a temperature from about 1200° F. to about 1400° F. (˜649 to 760° C.). The majority of, and preferably substantially all of the hot regenerated catalyst particles can then be recycled to the FCC reaction zone where they contact injected FCC feed.

In some aspects related to low temperature, high conversion FCC processing, the regeneration process can be performed in an alternative manner. In such alternative aspects, a low value fuel stream can be used to provide fuel for the regenerator. This can remove the requirement that sufficient coke can be present on the catalyst during regeneration to achieve the desired regenerator temperature. Suitable alternative fuel sources for the regenerator can include methane, torch oil, and/or various refinery streams that have fuel value. As the reaction temperature in low temperature FCC processing can be lower, the regeneration process can be performed at a lower temperature. A regenerated catalyst temperature of about 550° C. to about 630° C., or about 550° C. to about 600° C., can be sufficient to maintain a FCC riser temperature of about 450° C. to about 482° C.

Additional Embodiments

Embodiment 1. A method for processing a paraffin-containing feed, comprising: passing a first portion of heat transfer particles from a burner of a reaction system to a reactor of the reaction system; passing a second portion of heat transfer particles from the burner to a side riser; exposing a feed comprising C₃₊ paraffins to the second portion of heat transfer particles in the side riser under paraffin to olefin conversion conditions to form a side riser effluent comprising heat transfer particles including deposited coke and a gas phase effluent; separating a third portion of the heat transfer particles including deposited coke from the gas phase effluent; exposing at least a portion of the gas phase effluent to one or more beds of a conversion catalyst to form an aromatic formation effluent comprising C₆-C₁₂ aromatics; passing the third portion of the heat transfer particles into the reaction system; exposing a second feedstock to the first portion of heat transfer particles and the third portion of the heat transfer particles in the reactor under first processing conditions to form a first reactor effluent comprising heat transfer particles including additional coke; separating at least a portion of the heat transfer particles including additional coke from the first reactor effluent; and passing the separated heat transfer particles into the burner.

Embodiment 2. The method of Embodiment 1, wherein the reactor comprises a fluidized coking reactor, the first processing conditions comprise fluidized coking conditions, and the heat transfer particles comprise coke particles.

Embodiment 3. The method of Embodiment 2, wherein the burner comprises a gasifier.

Embodiment 4. The method of Embodiment 3, a) wherein transferring the first portion of heat transfer particles from the burner to the reactor comprises transferring the first portion of heat transfer particles to a heater of the reaction system, and transferring the first portion of heat transfer particles from the heater to the reactor; b) wherein passing the separated heat transfer particles into the burner comprises passing the separated heat transfer particles into a heater of the reaction system, and transferring the separated heat transfer particles from the heater to the burner; or c) a combination of a) and b).

Embodiment 5. The method of any of Embodiments 2-4, wherein passing the third portion of the heat transfer particles into the reaction system comprises passing the third portion of the heat transfer particles into the reactor; or wherein passing the third portion of the heat transfer particles into the reaction system comprises passing the third portion of the heat transfer particles into the burner; or wherein passing the third portion of the heat transfer particles into the reaction system comprises passing the third portion of the heat transfer particles into a heater of the reaction system; or a combination thereof.

Embodiment 6. The method of Embodiment 1, wherein the reactor comprises a fluid catalytic cracking (FCC) reactor, the first processing conditions comprise FCC processing conditions, the burner comprises a regenerator, and the heat transfer particles comprise FCC catalyst particles, the reaction system optionally further comprising a catalyst cooler.

Embodiment 7. The method of any of the above embodiments, wherein the one or more beds of the conversion catalyst comprise moving catalyst beds of conversion catalyst in a moving bed reactor; or wherein the one or more beds of the conversion catalyst comprise fixed beds of conversion catalyst in one or more radial flow reactors, one or more axial flow reactors, or a combination thereof.

Embodiment 8. The method of any of the above embodiments, wherein the feed comprises 30 vol % to 70 vol % C₃₊ paraffins, or wherein the feed comprises 30 vol % to 70 vol % C₃-C₄ paraffins, or wherein the feed comprises 30 vol % to 70 vol % C₃-C₆ paraffins, the feed optionally further comprising ethane, methane, or a combination of ethane and methane.

Embodiment 9. The method of any of the above embodiments, wherein the at least a portion of the gas phase effluent is exposed to the conversion catalyst at a temperature of about 450° C. to about 650° C., a pressure in the one or more beds of conversion catalyst comprising at least about 200 kPa-a, and a WHSV of about 0.1 hr⁻¹ to about 4.0 hr⁻¹.

Embodiment 10. The method of any of the above embodiments, the method further comprising heating an intermediate aromatic formation effluent formed by exposure of the at least a portion of the gas phase effluent to a first catalyst bed of the one or more beds of conversion catalyst, the heating of the intermediate aromatic effluent being prior to exposure of the intermediate aromatic formation effluent to a second catalyst bed of the one or more beds of conversion catalyst.

Embodiment 11. The method of any of the above embodiments, wherein the conversion catalyst comprises ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, MCM-22, MCM-49, or a combination thereof, the conversion catalyst further comprising 0.1 wt % to 5.0 wt % of a metal from Groups 3-13 of the periodic table relative to a weight of the catalyst particles, the metal optionally comprising Ga, In, or a combination thereof

Embodiment 12. A reaction system for conversion of paraffins to aromatics, comprising: a first reaction system comprising a reactor and a burner, the reactor being in fluid communication with the burner for transfer of heat transfer particles; a side riser comprising one or more side riser inlets and a side riser outlet, the one or more side riser inlets being in fluid communication with a feed source and in fluid communication with the burner for transfer of heat transfer particles; a separation stage comprising a separation stage inlet, a separation stage solids outlet, and a separation stage gas outlet, the separation stage inlet being in fluid communication with the side riser outlet, the separation stage solids outlet being in fluid communication with the first reactor; and a second reaction system comprising one or more beds of conversion catalyst, the second reaction system being in fluid communication with the separation stage gas outlet, wherein the conversion catalyst comprises a) 0.1 wt % to 5.0 wt % of a metal from Groups 3-13 of the periodic table relative to a weight of the catalyst particles, the metal optionally comprising Ga, In, or a combination thereof, and b) ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, MCM-22, MCM-49, or a combination thereof

Embodiment 13. The reaction system of Embodiment 12, wherein the one or more beds of conversion catalyst comprise moving beds of conversion catalyst in one or more moving bed reactors; or wherein the second reaction system comprises one or more fixed bed radial flow reactors comprising the one or more beds of conversion catalyst, one or more fixed axial flow reactors comprising the one or more beds of conversion catalyst, or a combination thereof

Embodiment 14. The reaction system of Embodiment 12 or 13, wherein the reactor comprises a fluidized coking reactor, the burner comprises a gasifier, and the heat transfer particles comprise coke particles, the first reaction system optionally further comprising a heater, wherein at least a portion of the fluid communication between the fluidized coking reactor and the gasifier comprises indirect fluid communication via the heater.

Embodiment 15. A gas phase effluent or an aromatic formation effluent formed by the method of any of Embodiments 1-11 or the reaction system of any of Embodiments 12-14.

Additional Embodiment A. The reaction system of Embodiment 12 or 13, wherein the reactor comprises a fluid catalytic cracking reactor and the burner comprises a regenerator, the reaction system optionally further comprising a catalyst cooler.

While the present invention has been described and illustrated by reference to particular embodiments, those of ordinary skill in the art will appreciate that the invention lends itself to variations not necessarily illustrated herein. For this reason, then, reference should be made solely to the appended claims for purposes of determining the true scope of the present invention. 

1. A method for processing a paraffin-containing feed, comprising: passing a first portion of heat transfer particles from a burner of a reaction system to a reactor of the reaction system; passing a second portion of heat transfer particles from the burner to a side riser; exposing a feed comprising C₃₊ paraffins to the second portion of heat transfer particles in the side riser under paraffin to olefin conversion conditions to form a side riser effluent comprising heat transfer particles including deposited coke and a gas phase effluent; separating a third portion of the heat transfer particles including deposited coke from the gas phase effluent; exposing at least a portion of the gas phase effluent to one or more beds of a conversion catalyst to form an aromatic formation effluent comprising C₆-C₁₂ aromatics; passing the third portion of the heat transfer particles into the reaction system; exposing a second feedstock to the first portion of heat transfer particles and the third portion of the heat transfer particles in the reactor under first processing conditions to form a first reactor effluent comprising heat transfer particles including additional coke; separating at least a portion of the heat transfer particles including additional coke from the first reactor effluent; and passing the separated heat transfer particles into the burner.
 2. The method of claim 1, wherein the reactor comprises a fluidized coking reactor, the first processing conditions comprise fluidized coking conditions, and the heat transfer particles comprise coke particles.
 3. The method of claim 2, wherein passing the third portion of the heat transfer particles into the reaction system comprises passing the third portion of the heat transfer particles into the reactor; or wherein passing the third portion of the heat transfer particles into the reaction system comprises passing the third portion of the heat transfer particles into the burner; or a combination thereof.
 4. The method of claim 2, wherein the burner comprises a gasifier.
 5. The method of claim 4, wherein transferring the first portion of heat transfer particles from the burner to the reactor comprises transferring the first portion of heat transfer particles to a heater of the reaction system, and transferring the first portion of heat transfer particles from the heater to the reactor.
 6. The method of claim 4, wherein passing the separated heat transfer particles into the burner comprises passing the separated heat transfer particles into a heater of the reaction system, and transferring the separated heat transfer particles from the heater to the burner.
 7. The method of claim 4, wherein passing the third portion of the heat transfer particles into the reaction system comprises passing the third portion of the heat transfer particles into a heater of the reaction system.
 8. The method of claim 1, wherein the reactor comprises a fluid catalytic cracking (FCC) reactor, the first processing conditions comprise FCC processing conditions, the burner comprises a regenerator, and the heat transfer particles comprise FCC catalyst particles.
 9. The method of claim 8, wherein the reaction system further comprises a catalyst cooler.
 10. The method of claim 1, wherein the one or more beds of the conversion catalyst comprise moving catalyst beds of conversion catalyst in a moving bed reactor; or wherein the one or more beds of the conversion catalyst comprise fixed beds of conversion catalyst in one or more radial flow reactors, one or more axial flow reactors, or a combination thereof.
 11. The method of claim 1, wherein the feed comprises 30 vol % to 70 vol % C₃₊ paraffins, or wherein the feed comprises 30 vol % to 70 vol % C₃-C₄ paraffins, or wherein the feed comprises 30 vol % to 70 vol % C₃-C₆ paraffins.
 12. The method of claim 1, wherein the feed further comprises methane, ethane, or a combination thereof.
 13. The method of claim 1, wherein the at least a portion of the gas phase effluent is exposed to the conversion catalyst at a temperature of about 450° C. to about 650° C., a pressure in the one or more beds of conversion catalyst comprising at least about 200 kPa-a, and a WHSV of about 0.1 hr⁻¹ to about 4.0 hr⁻¹.
 14. The method of claim 1, the method further comprising heating an intermediate aromatic formation effluent formed by exposure of the at least a portion of the gas phase effluent to a first catalyst bed of the one or more beds of conversion catalyst, the heating of the intermediate aromatic effluent being prior to exposure of the intermediate aromatic formation effluent to a second catalyst bed of the one or more beds of conversion catalyst.
 15. The method of claim 1, wherein the conversion catalyst comprises ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, MCM-22, MCM-49, or a combination thereof.
 16. The method of claim 1, wherein the conversion catalyst comprises 0.1 wt % to 5.0 wt % of a metal from Groups 3-13 of the periodic table relative to a weight of the catalyst particles, the metal optionally comprising Ga, In, or a combination thereof.
 17. A method for processing a paraffin-containing feed, comprising: passing a first portion of coke particles from a gasifier of a fluidized coking system, a heater of a fluidized coking system, or a combination thereof to a reactor of the fluidized coking system; passing a second portion of coke particles from the gasifier, the heater, or a combination thereof to a side riser; exposing a feed comprising C₃₊ paraffins to the second portion of coke particles in the side riser under paraffin to olefin conversion conditions to form a side riser effluent comprising coke particles including deposited coke and a gas phase effluent; separating a third portion of the coke particles including deposited coke from the gas phase effluent; exposing at least a portion of the gas phase effluent to one or more beds of a conversion catalyst to form an aromatic formation effluent comprising C₆-C₁₂ aromatics; passing the third portion of the coke particles into the fluidized coking reactor; exposing a fluidized coking feedstock to the first portion of coke particles and the third portion of the coke particles in the reactor under fluidized coking conditions to form a coker effluent comprising coke particles including additional coke; separating at least a portion of the coke particles including additional coke from the coker effluent; and passing the separated coker particles into the gasifier.
 18. A reaction system for conversion of paraffins to aromatics, comprising: a first reaction system comprising a reactor and a burner, the reactor being in fluid communication with the burner for transfer of heat transfer particles; a side riser comprising one or more side riser inlets and a side riser outlet, the one or more side riser inlets being in fluid communication with a feed source and in fluid communication with the burner for transfer of heat transfer particles; a separation stage comprising a separation stage inlet, a separation stage solids outlet, and a separation stage gas outlet, the separation stage inlet being in fluid communication with the side riser outlet, the separation stage solids outlet being in fluid communication with the first reactor; and a second reaction system comprising one or more beds of conversion catalyst, the second reaction system being in fluid communication with the separation stage gas outlet, wherein the conversion catalyst comprises a) 0.1 wt % to 5.0 wt % of a metal from Groups 3-13 of the periodic table relative to a weight of the conversion catalyst, the metal optionally comprising Ga, In, or a combination thereof, and b) ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, MCM-22, MCM-49, or a combination thereof.
 19. The reaction system of claim 18, wherein the one or more beds of conversion catalyst comprise moving beds of conversion catalyst in one or more moving bed reactors; or wherein the second reaction system comprises one or more fixed bed radial flow reactors comprising the one or more beds of conversion catalyst, one or more fixed axial flow reactors comprising the one or more beds of conversion catalyst, or a combination thereof.
 20. The reaction system of claim 18, wherein the reactor comprises a fluidized coking reactor, the burner comprises a gasifier, and the heat transfer particles comprise coke particles.
 21. The reaction system of claim 20, wherein the first reaction system further comprises a heater, wherein at least a portion of the fluid communication between the fluidized coking reactor and the gasifier comprises indirect fluid communication via the heater.
 22. The reaction system of claim 18, wherein the reactor comprises a fluid catalytic cracking reactor and the burner comprises a regenerator; or wherein the reaction system further comprises a catalyst cooler; or a combination thereof. 